Process for preparing gamma-butyrolactone, butane-1,4-diol and tetrahydrofuran

ABSTRACT

A process for the production of at least one C 4  compound selected from 1,4-butanediol, gamma-butyrolactone and tetrahydrofuran, in which a solution of maleic anhydride in a high boiling ester is esterified with a C 1  to C 4  alcohol to form the corresponding maleate ester which is hydrogenated to form the at least one C 4  compound. The specified high boiling ester has a boiling point which is about 30 degrees C. higher than that of the di-(C 1 -C 4  alkyl)maleate.

This application is a 371 of PCT/GB98/03264 filed Nov. 2, 1998.

This invention relates to the production of butane-1,4-diol,γ-butyrolactone and tetrahydrofuran.

Butane-1,4-diol, together with variable amounts of γ-butyrolactone andtetrahydrofuran, can be produced by hydrogenolysis of diesters of maleicacid, fumaric acid and mixtures thereof. A major use of butane-1,4-diolis as a feedstock for the plastics industry, particularly for theproduction of polybutylene terephthalate. It is also used as anintermediate for the production of γ-butyrolactone and of the importantsolvent, tetrahydrofuran.

The maleate and fumarate diesters used as feedstock for the productionof butane-1,4-diol by such a hydrogenolysis route are convenientlyprepared from maleic anhydride, which is itself produced by vapour phaseoxidation of a hydrocarbon feedstock, such as benzene, mixed C₄ olefins,or n-butane, in the presence of a partial oxidation catalyst. In thepartial oxidation of benzene there is typically used a supportedvanadium pentoxide catalyst promoted with MoO₃ and possibly otherpromoters. The reaction temperature is from about 400° C. to about 455°C. and the reaction pressure is from about 1 bar to about 3 bar, whileabout 4 times the theoretical amount of air is used in order to stayoutside the explosive limits. The contact time is about 0.1 s. When thefeedstock is a mixed C₄ olefin feedstock, i.e. a mixed butenesfeedstock, then the partial oxidation catalyst may be vanadium pentoxidesupported on alumina. Typical reaction conditions include use of atemperature of from about 425° C. to about 485° C. and a pressure offrom about 1.70 bar to about 2.05 bar. The volume ratio of air tobutenes may be about 75:1 in order to stay below explosive limits.Alternatively it is possible, according to more modern practice, todesign the plant so that satisfactory safe operation can be achieved,despite the fact that the feed mixture of air and butenes is within theflammable limits. In the case of n-butane as feedstock, the catalyst istypically vanadium pentoxide and the reaction conditions include use ofa temperature of from about 350° C. to about 450° C. and a pressure offrom about 1 bar to about 3 bar. The air:n-butane volume ratio may beabout 20:1, even though this may be within the flammable limits. Onedesign of reactor for such partial oxidation reactions comprisesvertical tubes surrounded by a jacket through which a molten salt iscirculated in order to control the reaction temperature.

In each case a hot vaporous reaction mixture is recovered from the exitend of the reactor which comprises maleic anhydride vapour, watervapour, carbon oxides, oxygen, nitrogen, and other inert gases, besidesorganic impurities such as formic acid, acetic acid, acrylic acid, andunconverted hydrocarbon feedstock.

One way of recovering maleic anhydride from such a reaction mixture isto cool it to about 150° C. using a steam-producing stream and then tocool it further to about 60° C. by cooling it against water in order tocondense part of the maleic anhydride, typically about 30% to about 60%of the maleic anhydride present. The remainder of the stream is thenscrubbed with water.

Scrubbing with water or with an aqueous solution or slurry is described,for example, in U.S. Pat. No. 2,638,481. Such scrubbing results inproduction of a solution of maleic acid which is then dehydrated, bydistilling with xylene, for example, so as to remove the water andre-form the anhydride. A disadvantage of such a procedure, however, isthat an unacceptable proportion of the product remains in the vapourphase. In addition, some of the maleic acid is inevitably isomerised tofumaric acid. The byproduct fumaric acid represents a loss of valuablemaleic anhydride and is difficult to recover from the process systemsince it tends to form crystalline masses which give rise to processproblems.

Because of this isomerisation problem a variety of other anhydrousscrubbing liquids have been proposed. For example, dibutyl phthalate hasbeen proposed as scrubbing liquid in GB-A-727828, GB-A-763339, andGB-A-768551. Use of dibutyl phthalate containing up to 10 weight %phthalic anhydride is suggested in U.S. Pat. No. 4,118,403. U.S. Pat.No. 3,818,680 teaches use of a normally liquid intramolecular carboxylicacid anhydride, such as a branched chain C₁₂₋₁₅-alkenyl substitutedsuccinic anhydride, for absorption of maleic anhydride from the reactionmixture exiting the partial oxidation reactor. Tricresyl phosphate hasbeen proposed for this purpose in FR-A-1125014. Dimethyl terephthalateis suggested for this duty in JP-A-32-8408 and dibutyl maleate inJP-A-35-7460. A high molecular weight wax as scrubbing solvent is taughtin U.S. Pat. No. 3,040,059, while U.S. Pat. No. 2,893,924 proposesscrubbing with diphenylpentachloride. Use of an aromatic hydrocarbonsolvent having a molecular weight between 150 and 400 and a boilingpoint above 140° C. at a temperature above the dew point of water in thevaporous reaction mixture, for example dibenzylbenzene, is suggested inFR-A-2285386. Absorption of maleic anhydride from the vaporous partialoxidation reaction mixture in dimethylbenzophenone followed bydistillation is described in U.S. Pat. No. 3,850,758.Polymethylbenzophenones, at least a portion of which contain at least 3methyl groups, can be used as liquid absorbent for maleic anhydrideaccording to U.S. Pat. No. 4,071,540. Dialkyl phthalate esters having C₄to C₈ alkyl groups and a total of 10 to 14 carbon atoms in both alkylgroups are proposed for absorption of maleic anhydride from the reactionmixture in U.S. Pat. No. 3,891,680. An ester of a cycloaliphatic acid,for example dibutyl hexahydrophthalate, is suggested as absorptionsolvent for maleic anhydride in ZA-A-80/1247.

It has also been proposed to effect direct condensation of maleicanhydride from the reaction mixture exiting the partial oxidationreactor. However, this procedure is inefficient because an unacceptableproportion of the maleic anhydride remains in the vapour phase.

The maleic anhydride product recovered following condensation or byscrubbing or absorption and distillation is then reacted with a suitableC₁ to C₄ alkanol, such as methanol or ethanol, to yield thecorresponding di-(C₁ to C₄ alkyl maleate. This di-(C₁ to C₄ alkyl)maleate may contain a minor amount of the corresponding di-(C₁ to C₄alkyl) fumarate, besides traces of the corresponding mono-(C₁ to C₄alkyl) maleate and/or fumarate. It is then subjected to hydrogenolysisto yield a mixture of butane-1,4-diol, together with variable amounts ofγ-butyrolactone and tetrahydrofuran, depending upon the hydrogenolysisconditions that are selected, and of the C₁ to C₄ alkanol which can berecycled to produce further di-(C₁ to C₄ alkyl) maleate.

Processes and plant for the production of dialkyl maleates from maleicanhydride are described, for example, in U.S. Pat. No. 4,795,824 and inWO-A-90/08127. This last mentioned document describes a column reactorcontaining a plurality of esterification trays each having apredetermined liquid hold-up and containing a charge of a solidesterification catalyst, such as an ion exchange resin containingpendant sulphonic acid groups. A liquid phase containing, for example, acarboxylic acid component flows down the column from one esterificationtray to the next lower one against an upflowing stream of vapour of thelower boiling component of the esterification reagents, typically the C₁to C₄ alkanol. Water of esterification is removed from the top of thecolumn reactor in a vapour stream, while ester product is recovered fromthe sump of the reactor. As the liquid flows down the trays itencounters progressively drier reaction conditions and theesterification reaction is driven further towards 100% ester formation.This column reactor may be followed by a polishing reactor operatingunder liquid phase reaction conditions, the ester-containing stream fromthe bottom of the column reactor being admixed with further C₁ to C₄alkanol prior to admission to the polishing reactor. When used for theproduction of a di-(C₁ to C₄ alkyl) maleate, the column reactor can bepreceded by a non-catalytic monoesterification reactor in which maleicanhydride is reacted with the C₁ to C₄ alkanol in the absence of anadded catalyst to form the mono-(C₁ to C₄ alkyl) maleate.

The hydrogenation of dialkyl maleates to yield butane-1,4-diol isdiscussed further in U.S. Pat. No. 4,584,419, U.S. Pat. No. 4,751,334and WO-A-88/00937, the disclosures of all of which are hereinincorporated by reference.

It would be desirable to simplify the production of butane-1,4,-diol,γ-butyrolactone and tetrahydrofuran, from maleic anhydride by the di-(C₁to C₄ alkyl) maleate hydrogenolysis route. In particular it would bedesirable to reduce the capital cost of construction of such a plant andalso to reduce its running costs, thereby making butane-1,4-diol,γ-butyrolactone and tetrahydrofuran more readily available.

It is accordingly an object of the present invention to simplify theproduction of butane-1,4,-diol, γ-butyrolactone and tetrahydrofuran frommaleic anhydride by the di-(C₁ to C₄ alkyl) maleate hydrogenolysisroute. A further object is to reduce the capital cost of construction ofsuch a plant by reducing significantly the numbers of distillationcolumns and the amount of other equipment required. It further seeks toreduce the running costs of a butane-1,4-diol production plant, therebymaking butane-1,4-diol, γ-butyrolactone and tetrahydrofuran more readilyavailable.

According to the present invention there is provided a process for theproduction of at least one C₄ compound selected from butane-1,4-diol,γ-butyrolactone and tetrahydrofuran, which includes the step ofhydrogenation in the vapour phase of a di-(C₁ to C₄ alkyl) maleate inthe presence of a particulate ester hydrogenation catalyst, whichprocess comprises:

(a) contacting a vaporous stream containing maleic anhydride vapour,water vapour, and carbon oxides in an absorption zone with a highboiling ester as solvent thereby to form a solution of maleic anhydridein the high boiling ester, said high boiling ester having a boilingpoint at atmospheric pressure which is at least about 30° C. higher thanthat of the di-(C₁ to C₄ alkyl) maleate and being selected from di-(C₁to C₄ alkyl) esters of alkyl dicarboxylic acids containing up to 13carbon atoms, mono- and di-(C₁₀ to C₁₈ alkyl) esters of maleic acid,fumaric acid, succinic acid, and mixtures thereof, (C₁ to C₄ alkyl)esters of naphthalenemonocarboxylic acids, tri-(C₁ to C₄ alkyl) estersof aromatic tricarboxylic acids, and di-(C₁ to C₄ alkyl) esters ofisophthalic acid;

(b) recovering from the absorption zone a waste gas stream;

(c) reacting maleic anhydride in the solution of maleic anhydride ofstep (a) under esterification conditions in an esterification zone witha C₁ to C₄ alkanol to form the corresponding di-(C₁ to C₄ alkyl)maleate;

(d) recovering from the esterification zone a solution of the di-(C₁ toC₄ alkyl) maleate in the high boiling ester;

(e) contacting the solution of the di-(C₁ to C₄ alkyl) maleate in thehigh boiling ester with a gaseous stream containing hydrogen thereby tostrip di-(C₁ to C₄ alkyl) maleate therefrom and to form a vaporousstream comprising hydrogen and di-(C₁ to C₄ alkyl) maleate;

(f) contacting material of the vaporous stream of step (e) in ahydrogenation zone under ester hydrogenation conditions with aheterogeneous ester hydrogenation catalyst thereby to convert di-(C₁ toC₄ alkyl) maleate by hydrogenation to at least one C₄ compound selectedfrom butane-1,4-diol, γ-butyrolactone and tetrahydrofuran; and

(g) recovering from the hydrogenation zone a product stream containingsaid at least one C₄ compound.

Preferably in such a process the C₁ to C₄ alkanol is methanol or ethanoland the di-(C₁ to C₄ alkyl) maleate is dimethyl maleate or diethylmaleate. The use of methanol as the C₁ to C₄ alkanol and of dimethylmaleate as the di-(C₁ to C₄ alkyl) maleate is especially preferred.

The vaporous stream of step (a) of the process of the invention ispreferably produced by partial oxidation of a hydrocarbon feedstock inthe presence of a partial oxidation catalyst using molecular oxygen,typically in the form of air. The hydrocarbon feedstock can be benzene,or a mixed C₄ olefin stream, but is most preferably n-butane. The use ofn-butane as hydrocarbon feedstock is currently preferred upon thegrounds of cost since it is a cheaper feedstock than benzene or butenes.Hence in the process of the invention the feedstock used for productionof the maleic anhydride containing vaporous stream of step (a) is mostpreferably n-butane and the catalyst is preferably vanadium pentoxide.Typical partial oxidation conditions in this case include use of atemperature of from about 350° C. to about 450° C. and a pressure offrom about 1 bar to about 3 bar, an air to n-butane ratio of from about15:1 to about 50:1, e.g. about 20:1 and a partial oxidation catalystcomprising vanadium pentoxide; the contact time is typically from about0.01 s to about 0.5 s, e.g. about 0.1 s.

Partial oxidation of the hydrocarbon feedstock is conveniently conductedin a reactor which comprises vertical tubes surrounded by a jacketthrough which a molten salt is circulated in order to control thereaction temperature. The vaporous stream from the partial oxidationreactor can then be cooled by external cooling with boiler feed water toraise steam, and possibly also by further external cooling with coolingwater to a temperature in the range of from about 60° C. to about 160°C.

In step (a) of the process of the invention the vaporous maleicanhydride stream is preferably contacted with the high boiling ester ata temperature in the range of from about 60° C. to about 160° C.,preferably from about 80° C. to about 120° C., and at a pressure of fromabout 1 bar to about 3 bar so as to form a solution comprising maleicanhydride in the high boiling ester. The contacting can be carried outby bubbling the vaporous stream through a body of the high boilingester. Alternatively the high boiling ester can be sprayed into thevaporous stream. Countercurrent contacting devices can also be employedwherein the ascending vaporous stream is contacted by a descendingstream of high boiling ester in a gas-liquid contacting device, such asa packed scrubber tower or a scrubber tower provided with trays. In thisstep the high boiling ester will typically be at a lower temperaturethan the vaporous stream so that the latter is cooled.

In the resulting solution of maleic anhydride in the high boiling esterthe concentration of maleic anhydride in the high boiling ester mayrange from about 100 g/l to about 400 g/l.

The high boiling ester has a boiling point at mospheric pressure that isat least about 30° C. higher, and preferably at least about 60° C. toabout 70° C. higher, than that of the di-(C₁ to C₄ alkyl) maleate.

As examples of esters of alkyl dicarboxylic acids containing up to 13carbon atoms from which a suitable high boiling ester can be selectedthere can be mentioned dimethyl, diethyl, di-n- or -iso-propyl, di-n-,-sec-, or iso-butyl esters of suberic acid, azelaic acid, sebacic acid,undecanedioic acid, dodecanedioic acid, and tridecanedioic acid. It ispreferred that the alkyl moiety in such an ester shall be derived fromthe same alkanol as the C₁ to C₄ alkanol used in the esterification step(c). In this way any transesterification reactions that may occur do notgive rise to additional esters. Thus when the alkanol used is methanoland the dialkyl maleate is dimethyl maleate, any ester used as the highboiling ester is preferably also a dimethyl ester, such as dimethylsebacate.

The high boiling ester may alternatively be selected from mono- anddi-(C₁₀ to C₁₈ alkyl) esters of one of the C₄ alkyl dicarboxylic acids,i.e. maleic acid, fumaric acid, and succinic acid, and mixtures thereof.Examples of such esters include the esters and mixtures thereof derivedfrom n-decyl alcohol, lauryl alcohol, myristyl alcohol, cetyl alcohol,stearyl alcohol, and eicosanol. In this case some hydrolysis of the highboiling ester may occur in the esterification zone resulting inliberation of a minor proportion of the corresponding C₁₀ to C₁₈ alkylalcohol. In addition some transesterification may occur in theesterification zone resulting in formation of a minor amount of amono-(C₁ to C₄ alkyl) mono-(C₁₀ to C₁₈ alkyl) ester of the C₄ alkyldicarboxylic acid. For example, if dilauryl maleate is used as the highboiling ester and if methanol is used as the C₁ to C₄ alkanol, then aminor amount of methyl lauryl maleate can be formed bytransesterification. However, the formation of these minor byproducts isnot disadvantageous even if the high boiling ester used in step (a)comprises recycled material resulting from step (e) because any free C₁₀to C₁₈ alkanol can react with fresh maleic anhydride in step (a) to formfresh mono- or di-(C₁₀ to C₁₈ alkyl) maleate. In addition any mono-(C₁to C₄ alkyl) mono-(C₁₀ to C₁₈ alkyl) ester of the C₄ alkyl dicarboxylicacid can undergo transesterification on the next occasion that it passesthrough the esterification zone to form the desired solvent or thedesired di-(C₁ to C₄ alkyl) maleate.

The high boiling ester may alternatively be selected from (C₁ to C₄alkyl) esters of naphthalenemonocarboxylic acids, such as methylnaphthalene-2-carboxylate, from tri-(C₁ to C₄ alkyl) esters of aromatictricarboxylic acids, such as trimethyl benzene-1,2,4-tricarboxylate, orfrom di-(C₁ to C₄ alkyl) esters of isophthalic acid, such as dimethylisophthalate.

The high boiling ester used in step (a) conveniently comprises materialresulting from the hydrogen stripping step (e). Hence it may containalready some di-(C₁ to C₄ alkyl) maleate.

Provided that appropriate conditions are adopted in step (a), the gasstream recovered in step (b) of the process of the invention can beessentially free from maleic anhydride.

Esterification of the maleic anhydride with the C₁ to C₄ alkanol iseffected in step (c) in an esterification zone. This may comprise anon-catalytic reactor in which the maleic anhydride in the solution inthe high boiling ester undergoes reaction in the absence of addedcatalyst with the C₁ to C₄ alkanol to form the corresponding mono-(C₁ toC₄ alkyl) maleate. The reaction is:

where R is a C₁ to C₄ alkyl radical. Some conversion of the mono-(C₁ toC₄ alkyl) maleate to the corresponding di-(C₁ to C₄ alkyl) maleate mayalso occur. The reaction concerned is:

where R is as defined above.

Such a non-catalytic reactor can be operated under monoesterificationconditions which typically comprise use of a temperature of from about65° C. to about 260° C. and a pressure of from about 1 bar to about 50bar. This can be followed by a catalytic esterification stage. Forexample, the catalytic esterification stage may comprise a plurality ofstirred tank reactors such as is disclosed in U.S. Pat. No. 4,795,824.Preferably, however, the catalytic esterification stage comprises acolumn reactor of the type disclosed in WO-A-90/03127. In this case thenon-catalytic monoesterification stage may comprise a stirred tankreactor or a column reactor containing one or more trays which do notcontain any esterification catalyst and which is fed from the bottomwith methanol or other C₁ to C₄ alkanol vapour, while the maleicanhydride solution from step (a) is fed downward through the columnreactor.

If the catalytic esterification stage comprises a column reactor of thetype disclosed in WO-A-90/03127, then the solution of maleic anhydride(or a solution comprising the corresponding mono-(C₁ to C₄ alkyl)maleate, if a separate monoesterification stage is used) in the highboiling ester is fed to the top esterification tray of the columnreactor, while an excess of C₁ to C₄ alkanol vapour is fed to the bottomof the reactor.

In the column reactor the esterification trays each hold a charge of asolid esterification catalyst. Each tray has a vapour upcomer means topermit vapour to enter the tray from below and to agitate the mixture ofliquid and solid esterification catalyst in a zone of turbulence on thetray and to keep the catalyst particles in suspension. In order avoidthe danger of “hot spots” forming on the tray through formation ofpockets of settled catalyst particles, the floor of each tray ispreferably designed so as to slope towards the zone of turbulence at aslope which exceeds the angle of repose of the catalyst particles underthe liquid. In addition each esterification tray has a downcomer meanswhich permits liquid, but not catalyst particles, to flow down from thattray to the next lower one. Such a downcomer means will usually beprovided with a screen to prevent catalyst particles passing downwardlytherethrough.

Typical reaction conditions in the column reactor include use of atemperature and pressure under which the C₁ to C₄ alkanol distils. Suchtemperature and pressure conditions will vary in dependence upon the C₁to C₄ alkanol selected but will typically include use of a temperatureof from about 65° C. to about 135° C. and a pressure of from about 1 barto about 3 bar. A typical solid esterification catalyst is the ionexchange resin sold under the designation Amberlyst™ 16 by Rohm and Haas(U.K.) Limited of Lennig House, 2 Mason's Avenue, Croydon CR9 3NB,England or that available as DPT1 ion exchange resin from KvaernerProcess Technology Limited of The Technology Centre, Princeton Drive,Thornaby, Stockton-on-Tees TS17 6PY, England.

In passing up the column from one esterification tray to the next higherone, the upflowing C₁ to C₄ alkanol vapour carries with it water ofesterification. Thus the di-(C₁ to C₄ alkyl) maleate-containing liquidpassing down the column reactor from one esterification tray to the nextlower one encounters drier and drier conditions as it proceeds down thecolumn. In this way the esterification reaction leading to formation ofthe di-(C₁ to C₄ alkyl) maleate is driven further and further towards100% conversion to the di-(C₁ to C₄ alkyl) maleate.

Any byproduct acid, such as acetic acid or acrylic acid, that is alsopresent in the vaporous stream from the partial oxidation reactor,together with any maleic acid or fumaric acid present in the solutionsupplied to the esterification zone, will undergo conversion to thecorresponding C₁ to C₄ alkyl ester or diester, as the case may be.

The vapour phase stream emerging from the topmost esterification traycomprises C₁ to C₄ alkanol vapour and water vapour; it may furtherinclude traces of minor byproducts such as the di-(C₁ to C₄ alkyl)ether, besides traces of the di-(C₁ to C₄ alkyl) maleate and of the C₁to C₄ alkyl acrylate. A further additional tray or trays may be providedabove the uppermost esterification tray to act as a form of washingcolumn in order to return di-(C₁ to C₄ alkyl) maleate to theesterification trays. The resulting vapour stream, which is nowessentially free from di-(C₁ to C₄ alkyl) maleate, exits the top of thecolumn.

From the bottom of the column reactor there is recovered a liquid streamcomprising a solution of the di-(C₁ to C₄ alkyl) maleate in the highboiling ester. This is essentially acid free. If desired this liquid canbe admixed with additional C₁ to C₄ alkanol and passed through apolishing reactor containing a bed of solid esterification catalystoperating under liquid phase operating conditions. Such conditionstypically include use of a temperature of from about 65° C. to about135° C. and a pressure of from about 1 bar to about 3 bar. A typicalsolid esterification catalyst is the ion exchange resin sold under thedesignation Amberlyst™ 16 by Rohm and Haas (U.K.) Limited of LennigHouse, 2 Mason's Avenue, Croydon CR9 3NB, England or that available asDPT1 ion exchange resin from Kvaerner Process Technology Limited of TheTechnology Centre, Princeton Drive, Thornaby, Stockton-on-Tees TS17 PY,England.

In step (e) of the process of the invention, a gas stream comprisinghydrogen is passed through the solution of the di-(C₁ to C₄ alkyl)maleate.

The hydrogen stripping step is preferably conducted substantially at orat a pressure slightly higher than the inlet pressure to the esterhydrogenation zone. The hydrogen stripping step is similarly preferablyconducted at substantially the desired inlet temperature to thehydrogenation step or a little below this temperature, for example fromabout 5° C. to about 20° C. below this temperature. Then the temperaturecan be raised to the desired inlet temperature by admixture of furtherhot hydrogen-containing gas which has the additional benefit of dilutingthe vaporous ester-containing stream and thereby ensuring that it is ata temperature above its dew point, preferably at least about 5° C.higher than its dew point.

The hydrogenation step is advantageously conducted in the vapour phase,using a heterogeneous ester hydrogenation catalyst. Typical esterhydrogenation catalysts include reduced promoted copper catalysts, forexample reduced copper chromite catalysts such as that sold under thedesignation PG 85/1 by Kvaerner Process Technology Limited of 20Eastbourne Terrace, London W2 6LE.

The catalyst particles preferably have a particle size in the range offrom about 0.5 mm to about 5 mm. The particles may be of any convenientshape, e.g. spheres, pellets, rings or saddles. When using a fixed bedof catalyst the reactor can be a shell-and-tube reactor, which can beoperated substantially isothermally; however, it is preferably anadiabatic reactor. The use of an adiabatic reactor is advantageous sinceits capital cost is much lower than that of a shell-and-tube reactor andit is generally much easier to charge the reactor with the chosencatalyst.

Hydrogenation is conducted at an elevated temperature of, for example,from about 150° C. to about 240° C. and at a pressure of from about 5bar to about 100 bar, preferably from about 50 bar to about 70 bar.

From the hydrogenation zone there is recovered a hydrogenation productmixture which contains, in addition to the C₁ to C₄ alkanol, alsobutane-1,4-diol, and some tetrahydrofuran and γ-butyrolactone. Even ifthe primary product of interest is butane-1,4-diol, the presence ofthese minor amounts of tetrahydrofuran and γ-butyrolactone is notdisadvantageous since these compounds are important chemicals ofcommerce and it is accordingly economic to recover them in pure form. Ifdesired, γ-butyrolactone can be recycled to the hydrogenation zone toproduce additional butane-1,4-diol. In addition the hydrogenolysisproduct mixture will normally contain minor amounts of the correspondingdi-(C₁ to C₄ alkyl) succinate, n-butanol, the corresponding dialkylalkoxysuccinate, e.g. dimethyl methoxysuccinate if the C₁ to C₄ alkanolis methanol, and water.

For further details regarding hydrogenation of a di-(C₁ to C₄ alkyl)maleate and subsequent purification of the resultant crude hydrogenationproduct mixture, reference may be made to U.S. Pat. No. 4,584,419,WO-A-86/03189, WO-A-88/0937, U.S. Pat. No. 4,767,869, U.S. Pat. No.4,945,173, U.S. Pat. No. 4,919,765, U.S. Pat. No. 5,254,758, U.S. Pat.No. 5,310,954, and WO-A-91/01960.

In order that the invention may be clearly understood and readilycarried into effect a plant for the production of butane-1,4-diol, aswell as some γ-butyrolactone and tetrahydrofuran, using a preferredprocess in accordance with the present invention will now be described,by way of example only, with reference to the accompanying drawing whichis a flow diagram of the plant.

Referring to the drawing, n-butane is supplied in line 1 at a pressureof from 1 to 3 bar and at a temperature of 400° C. to a partialoxidation plant 2 which is also supplied with air in line 3. Partialoxidation plant 2 is of conventional design and includes a partialoxidation reactor comprising tubes packed with a partial oxidationcatalyst consisting of vanadium pentoxide packed into tubes providedwith a jacket through which molten salt can be circulated for thepurpose of temperature control. The partial oxidation reactor isoperated at an air:n-butane feed ratio of 20:1.

A hot vaporous partial oxidation product stream is cooled by externalcooling against boiler feed water to raise steam and then againstcooling water to reduce its temperature to 138° C. It is recovered fromplant 2 in line 4. This contains 2.9% w/w maleic anhydride, 5.8% w/wwater, 1.3% w/w carbon dioxide, 1.0% w/w carbon monoxide, 0.01 w/wacetic acid, 0.01 w/w acrylic acid, 15.7w/w oxygen, and the balanceessentially comprising nitrogen and other inert gases. It is fed to thebottom of a scrubbing tower 5, up which it passes against a downflowingspray of dimethyl sebacate which is supplied at a temperature of about68° C. from line 6. The scrubbed waste gas stream which contains 0.03%w/w maleic anhydride exits the top of scrubbing tower 5 in vent gas line7 and is passed to a waste gas burner.

From the bottom of scrubbing tower 5 there is recovered a liquid streamin line 8 which comprises a solution of approximately 22% w/w maleicanhydride and 0.04% w/w acrylic acid in dimethyl sebacate. This issupplied to the top of a column reactor 9 of the type described inWO-A-90/08127. This comprises a number of esterification trays mountedone above the other, each containing a charge of a solid esterificationcatalyst, such as Amberlyst™ 16 resin or DPT1 ion exchange resin, andeach having a vapour upcomer for upflowing vapour and a liquid downcomerto permit liquid to flow down the column from one esterification tray tothe next lower one. Methanol vapour is supplied to the bottom of columnreactor by way of line 10. Water of esterification is removed in thevapour stream exiting the column reactor in line 11. Column reactor 9 isoperated at a temperature of from about 110° C. to about 125° C. and ata pressure of from about 1 bar to about 3 bar. The residence time in thecolumn reactor is about 3 hours. Normally the temperature on the toptray will be somewhat higher (e.g. about 125° C.) than that on thelowermost tray (e.g. about 115° C.).

A solution containing about 250 g/l dimethyl maleate in dimethylsebacate is withdrawn from the bottom of column reactor 9 in line 12 andpumped to near the top of a stripping column 13 which is operated at atemperature of 170° C. and a pressure of 885 psia (61.02 bar). Column 13has a number of distillation trays above the point of injection of thedimethyl maleate solution into column 13 so as to reduce carryover ofthe high boiling ester dimethyl sebacate in the overhead stream fromcolumn 13. The solution of dimethyl maleate in dimethyl sebacate flowsdown stripping column 13 against an upflowing stream of hydrogen fromline 14. The stripped dimethyl sebacate is recycled from the bottom ofstripping column 13 by way of line 6 to the top of scrubbing tower 5.From the top of stripping column 13 there emerges in line 15 a nearsaturated vapour mixture stream comprising dimethyl maleate in hydrogen,with a hydrogen:dimethyl maleate molar ratio of about 320:1. This vapourmixture stream is at a temperature of from about 180° C. to about 195°C. and at a pressure of 62 bar. It is diluted with further hot hydrogenfrom line 16 at a temperature of from about 180° C. to about 195° C. toyield a vaporous stream with a hydrogen:dimethyl maleate molar ratio ofabout 350:1 and is at least about 5° C. above its dew point.

This vaporous mixture passes onwards in line 17 to hydrogenation plant18 which includes an adiabatic reactor packed with a reducedcopper-based catalyst, for example, a reduced copper chromite catalyst,and operated at an inlet temperature of 173° C., an inlet pressure of885 psia (61.02 bar), and an exit temperature of 190° C. The dimethylmaleate feed rate corresponds to a liquid hourly space velocity of 0.5h⁻¹. The plant also includes a purification section in which the crudehydrogenation product mixture is distilled in several stages to yieldpure butane-1,4-diol in line 19. Lines for separate recovery ofγ-butyrolactone and tetrahydrofuran are indicated at 20 and 21respectively. Fresh dimethyl sebacate solvent can be added by means ofline 22 while a purge stream of the recycled solvent stream can be takenin line 23.

The dimethyl sebacate used as solvent can be replaced by, for example,methyl naphthalene-2-carboxylate, trimethyl benzene-1,2,4-tricarboxylateor dimethyl isophthalate. Alternatively the dimethyl sebacate can bereplaced as high boiling ester by a di-(C₁₀ to C₁₈ alkyl) maleate,fumarate, or succinate or a mixture of two or more thereof, optionallyin admixture with the corresponding mono-(C₁₀ to C₁₈ alkyl) maleate,fumarate, or succinate or a mixture of two or more thereof, and/or withthe corresponding free acid or mixture of acids, i.e. maleic acid,fumaric acid, and/or succinic acid. Typically the high boiling ester insuch a case comprises predominantly the diester or diester mixture, withno more than a minor amount, typically less than about 5 molar % each,of the corresponding monoester or monoester mixture and/or of thecorresponding acid or mixture of acids. As an example of such a highboiling ester there can be mentioned dilauryl maleate, which may containminor amounts, preferably less than about 1 molar % each, and even morepreferably less than about 0.25 molar % each, of one or more of dilaurylfumarate, dilauryl succinate, monolauryl maleate, monolauryl fumarate,monolauryl succinate, maleic acid, fumaric acid, and succinic acid. Inaddition, as a result of transesterification in the column reactor 9,the recycle stream in line 6 can in this case contain also a significantamount of methyl lauryl maleate, for example, up to about 10 molar & ormore, typically no more than about 5 molar %, and often less than about1 molar %; in addition it may contain minor amounts, typically less thanabout 1 molar % each, and even more preferably less than about 0.25molar % each, of lauryl alcohol, methyl lauryl fumarate, and methyllauryl succinate.

The invention is further illustrated by reference to the followingExamples.

EXAMPLE 1

98.0 g of methanol and 32.0 g of maleic anhydride were allowed to reacttogether in a round bottomed flask. To the resultant mixture was added130 g of dimethyl sebacate and 26.0 g of DPT1 ion exchange resin. (DPT1ion exchange resin is a macroreticular ion exchange resin containingsulphonic acid groups and is available from Kvaerner Process TechnologyLimited, of The Technology Centre, Princeton Drive, Thornaby,Stockton-on-Tees TS17 6PY, England). The mixture was heated to 110° C.and dry methanol fed to the flask at a rate of 6 mole hr⁻¹. Anyunconverted methanol, together with byproduct water and dimethyl ether,was recovered overhead following condensation. The conversion ofmonomethyl maleate to dimethyl maleate was followed by withdrawingsamples periodically from the reaction flask and analysing such samplesfor acid content. The experiment was continued until the level of aciddropped to less than 0.5% by weight. The results are shown in Table 1below.

TABLE 1 Time Overheads (minutes) wt/wt % acid wt (g) wt % DME DME (g)  055.9 — — — 30 18.3   72.5 0.04 0.03 60 2.9 114 0.09 0.14 90 0.41  940.19 0.31 120  0.27 101 0.31 0.62 150  0.12 110 0.36 1.02 Note: DME =dimethyl ether

The final sample was also analysed using a capillary gas chromatographictechnique. This enabled the amount of hydrolysis of dimethyl sebacatethat had occurred to be determined. The results are given in Table 2.

TABLE 2 Component wt/wt % Methanol 4.7 Dimethyl sebacate 45.8 Monomethylsebacate 0.06 Sebacic acid 0.01 Monomethyl maleate 0.18 Monomethylfumarate 0.14 Maleic acid 0.02 Dimethyl maleate 46.9 Dimethyl fumarate1.91 Unknowns 0.9 Water 0.09

The results indicate that a small quantity of dimethyl sebacate washydrolysed to monomethyl sebacate and sebacic acid when employing drymethanol.

EXAMPLE 2

32.0 g of maleic anhydride and 98.0 g of dry methanol were reactedtogether and then heated to 110° C. in 130 g of dimethyl sebacate in thepresence of 26.0 g of DPT1 ion exchange resin according to the procedureof Example 1. Methanol containing 30 mol % of water was then fed to theflask in the manner described in Example 1 at a rate of 6 mole hr⁻¹until the system reached equilibrium. Then the amount of water in themethanol fed to the flask was reduced to 15 mol % and the system againbrought to equilibrium. The experiment was continued using methanolcontaining 5 mol % water and finally using dry methanol. The resultsobtained are set out in Table 3.

TABLE 3 Mol % Product Dimethyl Monomethyl Sebacic water acidity sebacatesebacate acid 30 9.76 43.4 5.0 0.1 15 4.48 52.0 2.7 0.1  5 1.27 42.6 0.5Trace  0 0.16 59.7 0.2 0.1 Note: Product acidity means percentage ofmonomethyl maleate.

What is claimed is:
 1. A process for the production of at least one C₄compound selected from butane-1,4-diol, γ-butyrolactone andtetrahydrofuran, which includes the step of hydrogenation in the vapourphase of a di-(C₁ to C₄ alkyl) maleate in the presence of a particulateester hydrogenation catalyst, which process comprises: (a) contacting avaporous stream containing maleic anhydride vapour, water vapour, andcarbon oxides in an absorption zone with a high boiling ester as solventthereby to form a solution of maleic anhydride in the high boilingester, said high boiling ester having a boiling point at atmosphericpressure which is at least about 30° C. higher than that of the di-(C₁to C₄ alkyl) maleate and being selected from di-(C₁ to C₄ alkyl) estersof alkyl dicarboxylic acids containing up to 13 carbon atoms, mono- anddi-(C₁₀ to C₁₈ alkyl) esters of maleic acid, fumaric acid, succinicacid, and mixtures thereof, (C₁ to C₄ alkyl) esters ofnaphthalenemonocarboxylic acids, tri-(C₁ to C₄ alkyl) esters of aromatictricarboxylic acids, and di-(C₁ to C₄ alkyl) esters of isophthalic acid;(b) recovering from the absorption zone a waste gas stream; (c) reactingmaleic anhydride in the solution of maleic anhydride of step (a) underesterification conditions in an esterification zone with a C₁ to C₄alkanol to form the corresponding di-(C₁ to C₄ alkyl) maleate; (d)recovering from the esterification zone a solution of the di-(C₁ to C₄alkyl) maleate in the high boiling ester; (e) contacting the solution ofthe di-(C₁ to C₄ alkyl) maleate in the high boiling ester with a gaseousstream containing hydrogen thereby to strip di-(C₁ to C₄ alkyl) maleatetherefrom and to form a vaporous stream comprising hydrogen and di-(C₁to C₄ alkyl) maleate; (f) contacting material of the vaporous stream ofstep (e) in a hydrogenation zone under ester hydrogenation conditionswith a heterogeneous ester hydrogenation catalyst thereby to convertdi-(C₁ to C₄ alkyl) maleate by hydrogenation to at least one C₄ compoundselected from butane-1,4-diol, γ-butyrolactone and tetrahydrofuran; and(g) recovering from the hydrogenation zone a product stream containingsaid at least one C₄ compound.
 2. A process according to claim 1, inwhich the C₁ to C₄ alkanol is methanol and the di-(C₁ to C₄ alkyl)maleate is dimethyl maleate.
 3. A process according to claim 1, in whichthe vaporous stream of step (a) is produced by partial oxidation of ahydrocarbon feedstock in the presence of a partial oxidation catalystusing molecular oxygen.
 4. A process according to claim 3, in which thehydrocarbon feedstock is n-butane.
 5. A process according to claim 4, inwhich the partial oxidation catalyst comprises vanadium pentoxide and inwhich the partial oxidation conditions include use of a temperature offrom about 350° C. to about 450° C., a pressure of from about 1 bar toabout 3 bar, an air to n-butane ratio of from about 15:1 to about 50:1and a contact time of from about 0.01 s to about 0.5 s.
 6. A processaccording to claim 1, in which in step (a) the vaporous maleic anhydridestream is contacted with the high boiling ester at a temperature in therange of from about 60° C. to about 160° C. and at a pressure of fromabout 1 bar to about 3 bar so as to form a solution comprising maleicanhydride in the high boiling ester.
 7. A process according to claim 6,in which the contacting step is carried out in a countercurrentcontacting device wherein the ascending vaporous stream is contacted bya descending stream of solvent in a gas-liquid contacting device.
 8. Aprocess according to claim 1, in which the solvent is an alkyl esterwhose alkyl moiety is derived from the same alkanol as the C₁ to C₄alkanol used in the esterification step (c).
 9. A process according toclaim 1, in which the C₁ to C₄ alkanol is methanol, the di-(C₁ to C₄alkyl) maleate is dimethyl maleate, and the high boiling solvent is alsoa methyl ester.
 10. A process according to claim 9, in which the methylester is dimethyl sebacate.
 11. A process according to claim 9, in whichthe high boiling ester is naphthalene-2-carboxylate, trimethylbenzene-1,2,4-tricarboxylate, or dimethyl isophthalate.
 12. A processaccording to claim 1, in which the high boiling ester comprises adi-(C₁₀ to C₁₈ alkyl) ester of maleic acid, fumaric acid, succinic acid,or a mixture of two or more thereof.
 13. A process according to claim12, in which the high boiling ester further comprises a mono-(C₁₀ to C₁₈alkyl) ester of maleic acid, fumaric acid, succinic acid, or a mixtureof two or more thereof.
 14. A process according to claim 1, in which thehigh boiling ester used in step (a) comprises recycled materialresulting from the hydrogen stripping step (e).
 15. A process accordingto claim 1, in which the esterification zone comprises a non-catalyticreactor in which the maleic anhydride in the solution in the highboiling ester undergoes reaction in the absence of added catalyst withthe C₁ to C₄ alkanol to form the corresponding mono-(C₁ to C₄ alkyl)maleate.
 16. A process according to claim 1, in which the catalyticesterification stage comprises a column reactor provided with aplurality of esterification trays each of which holds a charge of asolid esterification catalyst, has a vapour upcomer means to permitvapour to enter the tray from below and to agitate the mixture of liquidand solid esterification catalyst in a zone of turbulence on the trayand to keep the catalyst particles in suspension, and a downcomer meanswhich permits liquid, but not catalyst particles, to flow down from thattray to the next lower one, the column reactor being supplied beneaththe lowermost esterification tray with a stream of C₁ to C₄ alkanolvapour and to an upper esterification tray with a solution in the highboiling ester comprising a material selected from maleic anhydride, amono-(C₁ to C₄ alkyl) maleate wherein the C₁ to C₄ alkyl group isderived from the C₁ to C₄ alkanol, and a mixture thereof.
 17. A processaccording to claim 16, in which the floor of each tray slopes towardsthe zone of turbulence at a slope which exceeds the angle of repose ofthe catalyst particles under the liquid.
 18. A process according toclaim 1, in which the esterification zone comprises an autocatalyticesterification zone wherein the esterification conditions include use ofa temperature of from about 70° C. to about 250° C., a pressure of fromabout 1 bar to about 50 bar and wherein maleic anhydride is converted byreaction with C₁ to C₄ alkanol at least in part to the correspondingmono-(C₁ to C₄ alkyl) maleate.
 19. A process according to claim 1,wherein the esterification zone includes a catalytic esterification zonewherein the esterification conditions include use of a temperature offrom about 65° C. to about 135° C. and of a solid esterificationcatalyst comprising an ion exchange resin containing pendant sulphonicacid groups.
 20. A process according to claim 1, in which the hydrogenstripping step is conducted at substantially the inlet pressure to theester hydrogenation zone.
 21. A process according to claim 1, in whichthe hydrogen stripping step is conducted at a temperature in the rangeof from the inlet temperature to the hydrogenation zone to about 20° C.below the inlet temperature to the hydrogenation zone.
 22. A processaccording to claim 1, in which the hydrogenation step is conducted inthe vapour phase using a reduced promoted copper catalyst at atemperature of from about 150° C. to about 240° C. and at a pressure offrom about 5 bar to about 100 bar.
 23. A process according to claim 1,in which there is recovered from the hydrogenation zone a hydrogenationproduct mixture which contains, in addition to butane-1,4-diol and theC₁ to C₄ alkanol, also minor amounts of tetrahydrofuran andγ-butyrolactone.
 24. A process according to claim 23, in which thehydrogenation product mixture is purified by distillation in one or morestages, including distillation in a “light ends” column to separateoverhead the volatile components of the mixture includingtetrahydrofuran, the C₁ to C₄ alkanol, water, and n-butanol.
 25. Aprocess according to claim 24, in which the bottoms product from the“light ends” column is further purified by distillation in one or morestages to yield pure butane-1,4-diol.